JP3744698B2

JP3744698B2 - Process for hydrogenating polycyclic aromatic hydrocarbons

Info

Publication number: JP3744698B2
Application number: JP28282198A
Authority: JP
Inventors: 重人畑中; 修定兼; 英壱岐
Original assignee: Nippon Oil Corp
Current assignee: Eneos Corp
Priority date: 1998-10-05
Filing date: 1998-10-05
Publication date: 2006-02-15
Anticipated expiration: 2018-10-05
Also published as: JP2000109853A

Description

【０００１】
【発明の属する技術分野】
本発明は、多環芳香族炭化水素を含有する石油留分を水素化処理する際に特定の水素化処理工程の組み合わせ、特定の触媒、かつ特定の水素化処理条件を用いることにより多環芳香族炭化水素を効率よく水素化する方法に関する。
【０００２】
【従来の技術】
原油の蒸留によって得られる直留軽油や減圧軽油あるいは重油の分解によって得られる分解軽油等は多環芳香族炭化水素を含んでおり、その量は直留軽油で１０重量％程度である。多環芳香族炭化水素としては２環の（アルキル）ナフタレン、３環の（アルキル）アントラセン、（アルキル）フェナントレン等が挙げられる。多環芳香族炭化水素を多く含む軽油をディーゼル燃料として使用するとパティキュレートの排出量が増加し環境を汚染する。そのため通常これら軽油は水素化処理され多環芳香族炭化水素の一部を水素化した後に燃料として使用される。しかしディーゼルエンジンから排出されるパティキュレートをさらに減少する必要が叫ばれており、そのためにはディーゼル燃料油中の多環芳香族炭化水素の量をさらに低減する必要があるといわれているが、そのための具体的手段が見出されなかった。
【０００３】
従来、多環芳香族炭化水素の水素化処理にはアルミナ等の多孔質担体にコバルトあるいはニッケルとモリブデンあるいはタングステンを担持した触媒が使用されてきた。しかしながら、この従来触媒では３環以上の芳香族炭化水素が水素化され難く、十分水素化するためには接触時間を大きくしたり圧力を非常に高くしなければならず、装置の建設コストや運転コストが極めて大きくなるという問題点があった。例えば、フェナントレンは部分水素化されテトラヒドロフェナントレンあるいはオクタヒドロフェナントレンには容易に水素化されるが、芳香環の全てを水素化するのは難しい。パティキュレートを低減するためには完全に芳香環を水素化するのが望ましい。
【０００４】
軽油の水素化処理に関しては、硫黄分の水素化脱硫を中心に研究が進められ数々の提案がなされている。例えば、難脱硫性の硫黄化合物に対して脱硫活性を高める方法として、触媒の担体にリンやホウ素を含有させた触媒（特開昭５２−１３５０３号）やゼオライトを担体に加えた触媒（特開平７−１９７０３９）などが報告されている。これらの触媒にはブレンシュテッド酸点が存在し、（ジ）メチルジベンゾチオフェンのメチル基を異性化したりフェニル基を水素化する能力が高く、４−メチルジベンゾチオフェンや４, ６−ジメチルジベンゾチオフェンの脱硫に対して高い活性を示す。
【０００５】
しかしながら、上記の担体にリンやホウ素あるいはゼオライトを担体に加えた触媒は、アルキルベンゾチオフェン類や４−あるいは６−位置にアルキル置換基を持たないジベンゾチオフェン類、例えばジベンゾチオフェン、１−、２−または３−メチルベンゾチオフェン等に対する脱硫活性が、従来から使用されてきたアルミナ担体にコバルトとモリブデンを担持した触媒より劣る欠点がある（F. van Looijら, Applied Catalysis A: General 170, 1-12 (1998) ）。又、ブレンシュテッド酸点が存在するため、製品が着色しやすく、オレフィンを含む原料を使用する場合や３５０℃以上の高温で反応に用いた場合はチオールやスルフィドが生成して脱硫率が低下する欠点もある。さらにブレンシュテッド酸点でオレフィン成分が重合してコーク析出し、触媒の失活が速いという大きな問題点もある。上記触媒を用いた場合は特に、原料油にオレフィンが含まれていない場合でも、硫黄化合物が脱硫される場合にはオレフィンを生成するのでコークの析出の原因となる。このことは、チオフェンを通油した場合のコーキング速度がオレフィンや芳香族を通油した場合のコーキング速度の１０倍にも達することから理解できる(Catalysis Review, 24, (3), 343 (1982)) 。
【０００６】
上記の各触媒を用いても０. ０５重量％あるいはそれ以下のレベルまで脱硫するのは難しく、方法及び反応装置の面から深度脱硫を達成する研究もなされている。例えば、特開平７−１０２２６６には反応条件の異なる２段階の反応によって色相を悪化させることなく深度脱硫を行う方法が提案されており、また特開平５−３１１１７９には蒸留によって脱硫の容易な軽質留分と脱硫し難い重質留分に分離し、それぞれを別々に水添脱硫後それぞれの生成物を一体化して深度脱硫を達成する方法が提案されている。しかしながら、反応条件の異なる２段階の反応によって色相を悪化させることなく深度脱硫を行う方法は色相の改善には効果があるものの深度脱硫をさらに進める効果はほとんどなく、蒸留によって脱硫の容易な軽質留分と脱硫し難い重質留分に分離し、それぞれを別々に水添脱硫後それぞれの生成物を一体化して深度脱硫を達成する方法にしても脱硫し難い重質留分に対しては高温・高圧を要するなど問題点が多い。
いずれにしろ、従来技術はいずれも、優れた性状の軽油を得るために、硫黄含有化合物の水素化脱硫を目指すものであり、本発明のごとき、多環式芳香族炭化水素の水素化反応に着目して軽油の多環成分の水素化処理を行なうものでなく、更には、パーティキュレートの減少を図るために多環芳香族炭化水素の水素化を図るものではない。
【０００７】
【発明が解決しようとする課題】
本発明の目的は、前記従来の問題点を解決し、沸点２００〜４００℃の石油留分中に含まれる多環芳香族炭化水素を効率的に水素化する方法を提供することにある。
【０００８】
【課題を解決するための手段】
本発明者らは前記の課題を解決するため、鋭意研究した結果、特定の工程の組合せ、装置構成、触媒、反応条件を用いることにより多環芳香族炭化水素を効率よく水素化する方法を見出し本発明を完成するに至った。
本発明は、第一に、多環芳香族炭化水素を含む沸点２００〜４００℃の石油系炭化水素の留分を水素化処理する際に、蒸留によって分離された軽質分中に含まれる３環以上の芳香族炭化水素が１重量％以下であり、かつ重質分中に含まれる２環芳香族炭化水素が１重量％以下となるように、蒸留により２環芳香族炭化水素を含む軽質分と３環以上の芳香族炭化水素を含む重質分とに分離し、軽質分と重質分をそれぞれ独立した工程で水素化処理した後に軽質分と重質分を混合することを含む多環芳香族炭化水素の水素化方法に関する。
第２に、本発明は、軽質分および重質分あるいはそれらのいずれか一方をそれぞれ独立した反応器で水素化処理した後に軽質分と重質分を混合することを含む、前記第１の発明の多環芳香族炭化水素の水素化方法に関する。
第３に、本発明は、蒸留により軽質分と重質分とを分離する際のカット温度が３３０〜３５０℃である、前記第１または第２の発明の多環芳香族炭化水素の水素化方法に関する。
第４に、本発明は、軽質分の水素化処理に際して、使用する触媒の担体が、アルミナ、アモルファスシリカアルミナ、ゼオライトのいずれかあるいはこれらの混合物であり、この担体上にコバルト及び／又はニッケルとモリブデンを担持した触媒を使用し、軽質分に対する水素化処理の反応条件が、温度３２０〜３８０℃、圧力５〜７ MPa 、 LHSV ０ . ５〜３ h ^-1 、水素 / 油比１０００〜３０００ scfb である、前記第１〜第３のいずれかの発明の多環芳香族炭化水素の水素化方法に関する。
第５に、本発明は、重質分中の硫黄含有量を５００ wtppm 以下とした後に、多孔質担体の上に白金、ロジウムまたはパラジウムを担持した触媒を使用し、重質分に対する水素化処理の反応条件が、温度２００〜３００℃、圧力１〜５ MPa 、 LHSV ０ . ５〜３ h ^-1 、水素 / 油比１０００〜５０００ scfb である、前記第１〜第４のいずれかの発明の多環芳香族炭化水素の水素化方法に関する。
第６に、本発明は、重質分の水素化処理に際して、使用する触媒の担体が、アルミナ、アモルファスシリカアルミナ、ゼオライトのいずれかあるいはこれらの混合物であり、この担体上にコバルト及び／又はニッケルとモリブデンを担持した触媒を使用し、重質分に対する水素化処理の反応条件が、温度３２０〜３６０℃、圧力７〜１５ MPa 、 LHSV ０ . ５〜３ h ^-1 、水素 / 油比２０００〜５０００ scfb である、前記第１〜第４のいずれかの発明の多環芳香族炭化水素の水素化方法に関する。
又、本発明の課題を解決するための手段に関して、本発明の具体的態様とその作用を以下に詳述する。
【０００９】
【発明の実施の形態と作用】
本発明は、水素化処理に当たってまず水素化処理しようとする多環芳香族炭化水素を含む沸点２００〜４００℃の石油留分を蒸留により２環芳香族炭化水素を含む軽質分と３環以上の芳香族炭化水素を含む重質分とに分離し、軽質分と重質分をそれぞれ独立した反応器で水素化処理した後に軽質分と重質分を混合して製品化する点に特徴を有する。蒸留による軽質分と重質分とのカット温度は、好ましくは３００〜３５０℃であり、さらに好ましくは３３０〜３５０℃である。これはアントラセンの沸点が３４２℃であり、フェナントレンの沸点が３３９℃であるため、３３０〜３５０℃でカットすれば水素化されにくい３環以上の芳香族炭化水素を重質分として、水素化されやすい２環芳香族炭化水素を含む軽質分から、分離できるからである。これを組成の面からみると、蒸留によって分離された軽質分中に含まれる３環以上の芳香族炭化水素が１重量％以下であり、かつ重質分中に含まれる２環芳香族炭化水素が１重量％以下となるように蒸留分離すると、その後の水素化反応を軽質分と重質分で独立して反応条件を設定できる本発明の特徴を十分発揮でき、高い水素化率を達成することができる。蒸留は通常の常圧多段式連続蒸留装置が使用でき、２環芳香族炭化水素を含む軽質分と３環以上の芳香族炭化水素を含む重質分との分離性能を勘案して蒸留段数を設定する。
【００１０】
蒸留で分離された２環芳香族炭化水素を含む軽質分は比較的緩やかな反応条件で水素化処理することができる。触媒としては、例えばアルミナ担体にコバルト及び／又はニッケルとモリブデンを担持した触媒を使用し、温度３２０〜３８０℃、圧力３〜２０ＭＰａ、ＬＨＳＶ０. ５〜３ｈ^-1、水素/ 油比１０００〜５０００ｓｃｆｂの条件の中から目標とする水素化率により設定できる。さらに好ましくは、温度３２０〜３８０℃、圧力５〜７ＭＰａ、ＬＨＳＶ０．５〜３ｈ^-1、水素/ 油比１０００〜３０００ｓｃｆｂの条件である。
【００１１】
蒸留で分離された３環以上の環芳香族炭化水素を含む重質分は軽質分の水素化反応よりも高圧で水素化処理する必要がある。
触媒としては、アルミナを主成分とする多孔質担体にコバルト及び／又はニッケルとモリブデンを担持した触媒を使用してもよいが、さらに水素化活性の高い触媒、例えばアルミナ８５〜９９重量％とゼオライト１〜１５重量％とを含む担体にニッケルとモリブデンを担持した触媒やアルミナまたはアモルファスシリカアルミナ８５〜９９重量％とゼオライト１〜２０重量％とを含む担体にニッケルとタングステンを担持した触媒を使用して水素化率を高めることもできる。これらの触媒の場合反応条件は、温度３２０〜３８０℃、圧力３〜２０ＭＰａ、ＬＨＳＶ０. ５〜３ｈ^-1、水素/ 油比１０００〜５０００ｓｃｆｂの条件の中から目標とする水素化率により設定できるが、軽質分の水素化よりも高圧、高接触時間の条件を設定する必要がある。さらに好ましい反応条件は、温度３２０〜３６０℃、圧力７〜１５ＭＰａ、ＬＨＳＶ０. ５〜３ｈ^-1、水素/ 油比２０００〜５０００ｓｃｆｂである。また、圧力を１０ＭＰａ以上とすると色相にも大変すぐれた生成油が得られる。
高圧装置は大型化するのが難しく、大型化のためにはコストの増加が避けられない。本発明では、水素化が難しい３環以上の芳香族炭化水素を分離した後に高圧下で水素化するので、全留分を高圧処理する場合に比べて装置の建設コストが極めて節約できる特徴がある。
【００１２】
また、触媒として多孔質担体の上に白金、ロジウムまたはパラジウム等の貴金属を担持した触媒を使用することもできる。貴金属触媒を使用する場合は、この条件よりさらに低温、低圧で反応することができる。これらの貴金属を担持した触媒を使用する場合、好ましい重質分に対する水素化処理の反応条件は、温度２００〜３００℃、圧力１〜５ＭＰａ、ＬＨＳＶ０. ５〜３ｈ^-1、水素/ 油比１０００〜５０００ｓｃｆｂである。
ただし、貴金属触媒は硫黄の被毒を受けるので、硫黄分を含む原料油を処理する場合には、脱硫反応器を前段に設置する必要がある。
【００１３】
反応温度については、芳香族の水素化反応に関しては、高温ほど反応の平行の面から水素化反応には不利となるため極端な高温での反応は避けるべきである。重質分の水素化反応は進行し難いので、必要に応じて並列または直列に配置した複数の反応器を使用してもよい。この場合、１段目の反応器出口水素ガス中に含まれている硫化水素をアミン吸収装置などにより分離するか、一部水素を系外に抜き出して代わりに新しい水素を導入するかして硫化水素濃度を低下させると２段めでの水素化反応が容易となる。１段目の反応で５００ｐｐｍ以下に硫黄分が低減できれば、２段目の水素化工程で白金、パラジウム等の貴金属触媒を使用することもできる。軽質分に関しても、必要に応じて第二反応器を設置して水素化処理をさらに進めてもよい。
貴金属触媒は特に重質分の水素化に対して有効で、全留分の水素化反応に高価な貴金属触媒を使用する必要は必ずしもなく、特定の重質留分に使用するのみで十分水素化が進行するのは本発明の重要な特徴である。原料油の脱硫処理は、全留分を脱硫した後に蒸留して軟質分と重質分に分離し、重質分を貴金属触媒で処理しても良いし、蒸留後に重質分のみを脱硫してから貴金属触媒により水素化しても良い。
【００１４】
こうして十分に低芳香族化された重質分は軽質分と混合され製品化することができる。重質分は軽質分の混合は蒸留で分離した際の比率で混合してもよいし、必要に応じて混合割合を変えて製品の蒸留性状を調整することもできる。他の脱硫装置で製造された軽油と混合して製品化することも可能である。これら混合により製品化する際に必要に応じて、潤滑性向上剤、セタン価向上剤、清浄剤を配合できるのは当然のことである。
【００１５】
こうして十分に低芳香族化された軽質分と重質分は混合され製品化することができる。軽質分と重質分の混合は蒸留で分離した際の比率で混合してもよいし、必要に応じて混合割合を変えて製品の蒸留性状を調整することもできる。他の水素化装置で処理された軽油留分と混合して製品化することも可能である。これら混合により製品化する際に必要に応じて、潤滑性向上剤、セタン価向上剤、清浄剤を配合できるのは当然のことである。
【００１６】
本発明に使用する触媒に担持する活性金属量は、通常の軽油の水素化処理触媒に採用されている量を採用することができる。すなわち担体の重量を１００重量部として（ゼオライト含めた重量）、ＣｏあるいはＮｉは酸化物換算で１〜１０重量部、好ましくは３〜６重量部であり、Ｍｏは酸化物換算で１０〜３０重量部、好ましくは１５〜２５重量部である。金属量は少ないと活性が不足し、また触媒の失活速度が大きくなる。一方、多すぎても活性が飽和してしまい不経済である。
ゼオライト１〜１５重量％を含む担体にニッケルとモリブデンを担持した触媒を使用する場合、ゼオライトとしてはＡ型ゼオライト、Ｘ型ゼオライト、Ｙ型ゼオライト、Ｌ型ゼオライト、ＭＦＩ型ゼオライト、モルデナイト等が使用できる。中でも、Ｙ型ゼオライトを脱アルミニウムして熱安定を高めたＵＳＹ型ゼオライトが最も好ましい。これらゼオライトはイオン交換してブレンシュテッド酸点を発現させるが、プロトン、アルカリ土類金属、希土類金属等でイオン交換することができる。
【００１７】
ゼオライトはアルミナのゲルに混合して成形して焼成してもよいし、成形したアルミナ担体にバインダーを用いて付着させてもよい。
各水素触媒の活性等を改善するために少量の各種改質成分を加えた触媒を使用してもさしつかえない。たとえば、燐を加えると金属の分散が良くなるとともにブレンシュテッド酸点が増加するため、水素化活性が向上する場合がある。一方、カリウムやマグネシウムの添加はブレンシュテッド酸点を減少させチオールやスルフィドの生成を抑制するので、各反応器の出口付近の触媒に加え、スウィートノング効果を期待することができる。
【００１８】
本発明が適用できる原料油は、直留（減圧）軽油、接触分解（減圧）軽油、熱分解（減圧）軽油等の沸点範囲が２００〜４００℃であり、かつ２環芳香族炭化水素および３環以上の芳香族炭化水素を含有するの石油留分である。本発明は沸点の４００℃よりもさらに高い減圧軽油の水素化処理にも有効である。
原料油に含まれる芳香族炭化水素の量は特に限定されないが、通常の直留軽油の場合は１０重量％程度である。生成油の芳香族炭化水素の量は必要に応じて任意に定めることができ、反応温度、圧力、液空間速度等の反応条件を最適化することにより必要とされる水素化率を達成できる。
また、硫黄化合物や窒素化合物を含む原料油を使用することもできる。この場合、芳香族の水素化と同時に水素化脱硫を達成して、低硫黄な製品を製造することができる。
本発明で水素化された軽油は、軽油自動車用レギュラーあるいはプレミアムディーゼル燃料として使用できる。また、Ａ重油等に混合して使用することもできる。
【００１９】
本発明の反応条件としては、温度３２０〜３８０℃、圧力３〜２０ＭＰａ、ＬＨＳＶ０. ５〜３ｈ^-1、水素/ 油比１０００〜５０００ｓｃｆｂの条件の中から目標とする水素化率により設定できる。
【００２０】
本発明に使用する反応器は、従来から知られているいかなる様式の反応器、例えば固定床、移動床いずれでも良く、ダウンフロー式、アップフロー式いずれでもよい。これらの中で最も適しているのは、固定床ダウンフロー式反応器である。これは従来から石油留分の水素化処理に用いられている反応器様式であるため、従来の装置をそのまま使用することができる。反応器は、通常軽質分および重質分用の各１反応器を複数の触媒ベットに分けたものが使用できる。本発明の水素化処理条件では液体と気体が共存するいわゆるトリクルベットであるため、各触媒ベットの上には液体を均一に分散させるディストリビューターを設置することが望ましい。また発熱状況により、クウェンチ水素を最適な場所で導入して発熱を制御してもよい。実際の装置には、押し出し成形した触媒が使用され、触媒は従来の方法によって反応器にソック充填またはデンス充填される。触媒を予備硫化した後、水素とともに加熱した原料油を触媒を充填した反応器に通油する。貴金属系の触媒は予備硫化せずに還元処理を行って使用する。使用済の触媒は通常の焼成再生処理によって繰り返し使用しても差し支えない。
【００２１】
【実施例】
本発明を実施例によりさらに詳細に説明する。
実施例１
中東系の直留軽油（沸点２３０〜３８２℃）を理論段数２０段の常圧蒸留装置で３４０℃をカット温度として軽質分７５容量％と重質分２５容量％に分離した。質量分析装置で芳香族炭化水素の割合を調べたところ、軽質分中の２環芳香族炭化水素の割合は９. ６重量％であり、３環芳香族炭化水素の割合は０. ３重量％であった。また、重質分中の２環芳香族炭化水素の割合は０. ６重量％であり、３環芳香族炭化水素の割合は４. ３重量％であり、４環芳香族炭化水素の割合は１. ２重量％であった。
第一反応器として、内径１インチの反応管にγ−アルミナ担体１００重量部に対してコバルト5 重量部（ＣｏＯ換算）とモリブデン２０重量部（ＭｏＯ₃換算）を担持した触媒を３００ｍｌ充填した。この触媒をジメチルジスルフィドを含む直留灯油（硫黄分３重量％）を用いて３００℃、５ＭＰａ、ＬＨＳＶ１ｈ^-1、水素/ 油比１０００ｓｃｆｂの条件下で、4 時間、予備硫化した後、軽質分を温度３５０℃、圧力５ＭＰａ、ＬＨＳＶ１ｈ^-1、水素/ 油比１０００ｓｃｆｂの条件で通油して水素化した。生成油の２環芳香族炭化水素の割合は３. ９重量％であり、３環芳香族炭化水素の割合は０. ２重量％であった。
さらに第二反応器として、内径１インチの反応管の上層部にγ−アルミナ９７重量％とプロトン交換ＵＳＹ型ゼオライト３重量％とを含む担体にニッケル３重量部（ＮｉＯ換算）とモリブデン２０重量部（ＭｏＯ₃換算）を担持した触媒を２００ｍｌ充填し、下層部にはγ−アルミナ担体にコバルト5 重量部（ＣｏＯ換算）とモリブデン２０重量部（ＭｏＯ₃換算）を担持した触媒を１００ｍｌ充填した。この触媒をジメチルジスルフィドを含む直留灯油（硫黄分３重量％）を用いて３００℃、５ＭＰａ、ＬＨＳＶ１ｈ^-1、水素/ 油比１０００ｓｃｆｂの条件下で、4 時間、予備硫化した後、前述の蒸留分離した重質分を温度３５０℃、圧力１０ＭＰａ、ＬＨＳＶ１ｈ^-1、水素/ 油比２０００ｓｃｆｂの条件で通油して水素化処理した。生成油の２環芳香族炭化水素の割合は０. ０重量％であり、３環芳香族炭化水素の割合は１. ３重量％であり、４環芳香族炭化水素の割合は０. ４重量％であった。この重質分と前述の軽質分を混合して、２環芳香族炭化水素の割合は２. ９重量％であり、３環芳香族炭化水素の割合は０. ５重量％であり、４環芳香族炭化水素の割合は０. １重量％、色はセイボルトカラー（ＪＩＳＫ−２５８０）で＋１５の軽油を製造した。
【００２２】
実施例２
実施例１の触媒に替えて、第一反応器および第二反応器にγ−アルミナ担体１００重量部に対してニッケル5 重量部（ＮｉＯ換算）とモリブデン２０重量部（ＭｏＯ₃換算）を担持した触媒をそれぞれ３００ｍｌづつ充填した。この触媒を実施例１と同様に予備硫化し、実施例１の軽油の軽質分を第一反応器で、重質分を第二反応器で実施例１と同一条件で脱硫した。第一反応器生成油の２環芳香族炭化水素の割合は３. ６重量％であり、３環芳香族炭化水素の割合は０. １重量％であり、第二反応器生成油の２環芳香族炭化水素の割合は０. ０重量％であり、３環芳香族炭化水素の割合は１. ５重量％であり、４環芳香族炭化水素の割合は０. ５重量％であった。この重質分と前述の軽質分を混合して、２環芳香族炭化水素の割合は２. ７重量％であり、３環芳香族炭化水素の割合は０. ４重量％であり、４環芳香族炭化水素の割合は０. １重量％、色はセイボルトカラーで＋２０の軽油を製造した。
【００２３】
実施例３
中東系の直留軽油（沸点２２４〜３６８℃）８０容量％と、接触分解軽油（沸点２１２〜３４５℃）１０容量％と、直脱分解軽油（沸点１８１〜３４６℃）１０容量％とを混合した。を理論段数２０段の常圧蒸留装置で３３５℃をカット温度として軽質分７７容量％と重質分２３容量％に分離した。軽質分中の２環芳香族炭化水素の割合は１１. ９重量％であり、３環芳香族炭化水素の割合は０. ４重量％であった。また、重質分中の２環芳香族炭化水素の割合は０. ７重量％であり、３環芳香族炭化水素の割合は５. ８重量％であり、４環芳香族炭化水素の割合は１. １重量％であった。
この混合軽油を実施例１と同一の触媒を同量充填した第一反応器で軽質分を、実施例１と同一の触媒を同量充填した第二反応器で重質分をそれぞれ実施例１と同一の反応条件で水素化した。第一反応器生成油の２環芳香族炭化水素の割合は４. ２重量％であり、３環芳香族炭化水素の割合は０. ２重量％であり、第二反応器生成油の２環芳香族炭化水素の割合は０. ０重量％であり、３環芳香族炭化水素の割合は１. ５重量％であり、４環芳香族炭化水素の割合は０. ４重量％であった。この重質分と前述の軽質分を混合して、２環芳香族炭化水素の割合は３. ２重量％であり、３環芳香族炭化水素の割合は０. ５重量％であり、４環芳香族炭化水素の割合は０. １重量％、色はセイボルトカラーで−２の軽油を製造した。
【００２４】
実施例４
水素化脱硫反応器として、内径１インチの反応管にγ−アルミナ担体１００重量部に対してコバルト5 重量部（ＣｏＯ換算）とモリブデン２０重量部（ＭｏＯ₃換算）を担持した触媒を３００ｍｌ充填した。この触媒を、ジメチルジスルフィドを含む直留灯油（硫黄分３重量％）を用いて温度３００℃、圧力５ＭＰａ、ＬＨＳＶ１ｈ^-1、水素/ 油比１０００ｓｃｆｂの条件下で、4 時間、予備硫化した。これに実施例１で用いた中東系直留軽油（沸点２３０〜３８２℃、硫黄分１．５重量％）を温度３５０℃、圧力５ＭＰａ、ＬＨＳＶ１ｈ^-1、水素/ 油比２０００ｓｃｆｂの条件下で通油して水素化脱硫処理した。生成油を理論段数２０段の常圧蒸留装置で３４０℃をカット温度として軽質分７７容量％と重質分２３容量％に分離した。質量分析装置で芳香族炭化水素の割合を調べたところ、軽質分中の２環芳香族炭化水素の割合は４．５重量％であり、３環芳香族炭化水素の割合は０. ２重量％であった。また、重質分中の２環芳香族炭化水素の割合は０．４重量％であり、３環芳香族炭化水素の割合は２. ３重量％であり、４環芳香族炭化水素の割合は０．８重量％であった。硫黄分は軽質分が０．０１重量％であり、重質分が０．０４重量％であった。
第一反応器として、内径１インチの反応管にγ−アルミナ担体１００重量部に対してニッケル5 重量部（Ｎi Ｏ換算）とモリブデン２０重量部（ＭｏＯ₃換算）を担持した触媒を３００ｍｌ充填した。この触媒を、ジメチルジスルフィドを含む直留灯油（硫黄分３重量％）を用いて、温度３００℃、圧力５ＭＰａ、ＬＨＳＶ１ｈ^-1、水素/ 油比１０００ｓｃｆｂの条件下で、4 時間、予備硫化した後、軟質分を温度３５０℃、圧力５ＭＰａ、ＬＨＳＶ１ｈ^-1、水素/ 油比１０００ｓｃｆｂの条件下で通油して水素化した。生成油の２環芳香族炭化水素の割合は２．３重量％であり、３環芳香族炭化水素の割合は０．１重量％であった。硫黄分は０．００１重量％であった。さらに第二反応器として、内径１インチの反応管にγ−アルミナ９７重量％とプロトン交換ＵＳＹ型ゼオライト３重量％とを含む担体にプラチナを０．５重量％担持した触媒を２００ｍｌ充填した。この触媒を３００℃、４時間水素還元した後、前述の蒸留分離した重質分を温度２５０℃、圧力２ＭＰａ、ＬＨＳＶ１ｈ^-1、水素/ 油比２０００ｓｃｆｂの条件下で通油して水素化処理した。生成油の２環芳香族炭化水素の割合は０．１重量％であり、３環芳香族炭化水素の割合は１．０重量％であり、４環芳香族炭化水素の割合は０．２重量％であった。硫黄分は０．００１重量％であった。
この重質分と前述の軽質分を混合して、２環芳香族炭化水素の割合は１．８重量％であり、３環芳香族炭化水素の割合は０. ３重量％であり、４環芳香族炭化水素の割合は０. １重量％、色はセイボルトカラー（ＪＩＳＫ−２５８０）で＋２１の軽油を製造した。硫黄分は０．００２重量％あった。
比較例１
内径１インチの反応管に反応管に、γ−アルミナ９７重量％とプロトン交換ＵＳＹ型ゼオライト３重量％とを含む担体にニッケル３重量部（ＮｉＯ換算）とモリブデン２０重量部（ＭｏＯ₃換算）を担持した触媒を３００ｍｌ充填した。ジメチルジスルフィドを含む直留灯油（硫黄分３重量％）を用いて３００℃、５ＭＰａ、ＬＨＳＶ１ｈ^-1、水素/ 油比１０００ｓｃｆｂの条件下で、4 時間、予備硫化した後、実施例で用いた軽油を温度３５０℃、圧力５ＭＰａ、ＬＨＳＶ１ｈ^-1、水素/ 油比１０００ｓｃｆｂの条件で通油して水素化した。生成油中の２環芳香族炭化水素の割合は３. ７重量％であり、３環芳香族炭化水素の割合は１. ０重量％であり、４環芳香族炭化水素の割合は０. ３重量％、色はセイボルトカラーで−１５であった。
【００２５】
比較例２
内径１インチの反応管に反応管に、γ−アルミナ担体１００重量部に対してコバルト5 重量部（ＣｏＯ換算）とモリブデン２０重量部（ＭｏＯ₃換算）を担持した触媒を６００ｍｌ充填した。この触媒を比較例１と同様に予備硫化し、実施例１の軽油を通油して比較例１と同一条件で脱硫した。生成油中の２環芳香族炭化水素の割合は３. ７重量％であり、３環芳香族炭化水素の割合は１. ２重量％であり、４環芳香族炭化水素の割合は０. ３重量％、色はセイボルトカラーで＋１５であった。
【００２６】
比較例３
比較例１の触媒に実施例３の混合軽油を通油して水素化脱硫した。反応条件は比較例１と同じである。生成油の２環芳香族炭化水素の割合は５. ２重量％であり、３環芳香族炭化水素の割合は２. ５重量％であり、４環芳香族炭化水素の割合は０. ５重量％、色はセイボルトカラーで−２０の軽油を製造した。
【００２７】
【発明の効果】
多環芳香族炭化水素を含む沸点２００〜４００℃の石油系炭化水素の留分を水素化処理する際に、本発明を採用することにより多環芳香族の含有率が低い、着色のほとんどない、優れた軽油が効率的に製造できる。[0001]
BACKGROUND OF THE INVENTION
The present invention provides a polycyclic aromatic by using a combination of specific hydroprocessing steps, a specific catalyst, and specific hydroprocessing conditions when hydrotreating a petroleum fraction containing polycyclic aromatic hydrocarbons. The present invention relates to a method for efficiently hydrogenating a group hydrocarbon.
[0002]
[Prior art]
A straight-run gas oil obtained by distillation of crude oil, a vacuum gas oil or a cracked gas oil obtained by cracking heavy oil contains polycyclic aromatic hydrocarbons, and the amount thereof is about 10% by weight in the straight-run gas oil. Examples of polycyclic aromatic hydrocarbons include bicyclic (alkyl) naphthalene, tricyclic (alkyl) anthracene, and (alkyl) phenanthrene. When diesel oil containing a large amount of polycyclic aromatic hydrocarbons is used as diesel fuel, particulate emissions increase and pollute the environment. Therefore, these light oils are usually hydrotreated and used as fuel after hydrogenating a part of polycyclic aromatic hydrocarbons. However, there is a need to further reduce the particulates emitted from diesel engines. To that end, it is said that the amount of polycyclic aromatic hydrocarbons in diesel fuel oil needs to be further reduced. No specific means was found.
[0003]
Conventionally, a catalyst in which cobalt or nickel and molybdenum or tungsten are supported on a porous carrier such as alumina has been used for the hydrotreatment of polycyclic aromatic hydrocarbons. However, with this conventional catalyst, aromatic hydrocarbons of 3 or more rings are difficult to be hydrogenated, and in order to fully hydrogenate, the contact time must be increased or the pressure must be very high. There was a problem that the cost was extremely high. For example, phenanthrene is partially hydrogenated and easily hydrogenated to tetrahydrophenanthrene or octahydrophenanthrene, but it is difficult to hydrogenate all of the aromatic rings. In order to reduce the particulates, it is desirable to completely hydrogenate the aromatic ring.
[0004]
With regard to hydroprocessing of diesel oil, research is proceeding mainly on hydrodesulfurization of sulfur content, and many proposals have been made. For example, as a method for increasing the desulfurization activity for a hardly desulfurizing sulfur compound, a catalyst in which phosphorus or boron is contained in a catalyst carrier (Japanese Patent Laid-Open No. 52-13503) or a catalyst in which zeolite is added to a carrier (Japanese Patent Laid-Open No. 7-197039) has been reported. These catalysts have Bronsted acid sites and have a high ability to isomerize the methyl group of (di) methyldibenzothiophene or to hydrogenate the phenyl group. 4-methyldibenzothiophene and 4,6-dimethyldibenzothiophene High activity against desulfurization.
[0005]
However, catalysts in which phosphorus, boron or zeolite is added to the above support are alkylbenzothiophenes or dibenzothiophenes having no alkyl substituent at the 4- or 6-position, such as dibenzothiophene, 1-, 2- Alternatively, the desulfurization activity for 3-methylbenzothiophene and the like is inferior to that of a catalyst in which cobalt and molybdenum are supported on a conventionally used alumina support (F. van Looij et al., Applied Catalysis A: General 170, 1-12). (1998)). In addition, because of the presence of Bronsted acid sites, products are easy to color, and when using raw materials containing olefins or when used in reactions at temperatures higher than 350 ° C, thiols and sulfides are produced and the desulfurization rate decreases. There are also disadvantages. In addition, the olefin component is polymerized at the Brönsted acid point and coke is precipitated, resulting in a serious problem that the deactivation of the catalyst is fast. In particular, when the above catalyst is used, even when the olefin is not contained in the raw material oil, when the sulfur compound is desulfurized, olefin is produced, which causes coke precipitation. This can be understood from the fact that the coking speed when thiophene is passed is 10 times the coking speed when olefin or aromatic is passed (Catalysis Review, 24, (3), 343 (1982) )
[0006]
Even if each of the above catalysts is used, it is difficult to desulfurize to a level of 0.05% by weight or less, and studies have been made to achieve deep desulfurization from the viewpoint of the method and the reactor. For example, Japanese Patent Laid-Open No. 7-102266 proposes a method of performing deep desulfurization without deteriorating the hue by a two-stage reaction with different reaction conditions, and Japanese Patent Laid-Open No. 5-3117979 is a light product that can be easily desulfurized by distillation. There has been proposed a method for achieving deep desulfurization by separating a distillate and a heavy fraction that is difficult to desulfurize, separately hydrodesulfurizing each, and then integrating the respective products. However, the method of performing deep desulfurization without deteriorating the hue by a two-stage reaction with different reaction conditions is effective in improving the hue, but has little effect of further deep desulfurization, and light distillation that is easy to desulfurize by distillation. And heavy fractions that are difficult to desulfurize, and each of them is hydrodesulfurized separately and then combined with each product to achieve deep desulfurization.・ There are many problems such as requiring high pressure.
In any case, all of the conventional techniques aim to hydrodesulfurize sulfur-containing compounds in order to obtain light oil with excellent properties, and in the hydrogenation reaction of polycyclic aromatic hydrocarbons as in the present invention. It does not focus on hydrogenation of polycyclic components of light oil, and further does not attempt to hydrogenate polycyclic aromatic hydrocarbons in order to reduce particulates.
[0007]
[Problems to be solved by the invention]
An object of the present invention is to solve the above-mentioned conventional problems and to provide a method for efficiently hydrogenating polycyclic aromatic hydrocarbons contained in a petroleum fraction having a boiling point of 200 to 400 ° C.
[0008]
[Means for Solving the Problems]
  As a result of diligent research to solve the above problems, the present inventors have found a method for efficiently hydrogenating polycyclic aromatic hydrocarbons by using specific process combinations, apparatus configurations, catalysts, and reaction conditions. The present invention has been completed.
  The present inventionFirst, when hydrotreating a fraction of petroleum hydrocarbons having a boiling point of 200 to 400 ° C. containing polycyclic aromatic hydrocarbons, three or more aromatics contained in the light fractions separated by distillation Distillation of light components including bicyclic aromatic hydrocarbons and three or more rings by distillation so that the hydrocarbons are 1 wt% or less and the bicyclic aromatic hydrocarbons contained in the heavy component are 1 wt% or less. Polycyclic aromatic hydrocarbons, which are separated into heavy components containing aromatic hydrocarbons, and the light components and the heavy components are hydrotreated in independent processes, followed by mixing the light components and the heavy components The present invention relates to a hydrogenation method.
    Second, the present invention includes mixing the light component and the heavy component after hydrotreating the light component and / or the heavy component in an independent reactor, respectively. The present invention relates to a method for hydrogenating polycyclic aromatic hydrocarbons.
    Third, the present invention provides the hydrogenation of the polycyclic aromatic hydrocarbon according to the first or second invention, wherein a cut temperature when separating a light component and a heavy component by distillation is 330 to 350 ° C. About the method
    Fourthly, in the present invention, in the hydrotreatment of the light component, the catalyst support used is any one of alumina, amorphous silica alumina, and zeolite, or a mixture thereof, and cobalt and / or nickel on the support. Using a catalyst supporting molybdenum, the reaction conditions of the hydrogenation treatment for the light component are a temperature of 320 to 380 ° C., a pressure of 5 to 7 MPa , LHSV 0 . 5-3 h ^-1 ,hydrogen / Oil ratio 1000-3000 scfb The method for hydrogenating a polycyclic aromatic hydrocarbon according to any one of the first to third inventions.
    Fifth, the present invention reduces the sulfur content in the heavy component to 500. wtppm After the following, using a catalyst supporting platinum, rhodium or palladium on a porous carrier, the reaction conditions for the hydrogenation treatment on the heavy component are 200 to 300 ° C., pressure 1 to 5 MPa , LHSV 0 . 5-3 h ^-1 ,hydrogen / Oil ratio 1000-5000 scfb The method for hydrogenating a polycyclic aromatic hydrocarbon according to any one of the first to fourth inventions.
    Sixth, according to the present invention, the catalyst support used in the hydrotreatment of the heavy component is any one of alumina, amorphous silica alumina and zeolite, or a mixture thereof, and cobalt and / or nickel is supported on the support. And a catalyst supporting molybdenum, the reaction conditions of the hydrogenation treatment for heavy components are as follows: temperature 320 to 360 ° C., pressure 7 to 15 MPa , LHSV 0 . 5-3 h ^-1 ,hydrogen / Oil ratio 2000-5000 scfb The method for hydrogenating a polycyclic aromatic hydrocarbon according to any one of the first to fourth inventions.
  Further, regarding the means for solving the problems of the present invention, specific embodiments and functions of the present invention will be described in detail below.
[0009]
BEST MODE FOR CARRYING OUT THE INVENTION
In the present invention, a petroleum fraction having a boiling point of 200 to 400 ° C. containing a polycyclic aromatic hydrocarbon to be hydrotreated first is subjected to a hydrotreating process by distilling a light fraction containing a bicyclic aromatic hydrocarbon and three or more rings. It is characterized in that it is separated into heavy components containing aromatic hydrocarbons, and the light components and heavy components are hydrotreated in separate reactors, and then the light components and heavy components are mixed to produce a product. . The cut temperature of the light and heavy components by distillation is preferably 300 to 350 ° C, more preferably 330 to 350 ° C. This is because anthracene has a boiling point of 342 ° C. and phenanthrene has a boiling point of 339 ° C., so that it is hydrogenated using heavy hydrocarbons of three or more aromatic hydrocarbons that are difficult to be hydrogenated if cut at 330 to 350 ° C. This is because it can be easily separated from light components containing bicyclic aromatic hydrocarbons. From the aspect of composition, the tricyclic or higher aromatic hydrocarbon contained in the light fraction separated by distillation is 1% by weight or less, and the bicyclic aromatic hydrocarbon contained in the heavy fraction. When it is distilled and separated so that the amount of water becomes 1% by weight or less, the characteristics of the present invention that can set the reaction conditions for the light and heavy components independently can be exhibited sufficiently, and a high hydrogenation rate is achieved. be able to. Distillation can be performed using a normal atmospheric multi-stage continuous distillation system, and the number of distillation stages can be determined in consideration of the separation performance of light components containing bicyclic aromatic hydrocarbons and heavy components containing three or more aromatic hydrocarbons. Set.
[0010]
Light components containing bicyclic aromatic hydrocarbons separated by distillation can be hydrotreated under relatively mild reaction conditions. As the catalyst, for example, a catalyst in which cobalt and / or nickel and molybdenum are supported on an alumina carrier is used, temperature 320 to 380 ° C., pressure 3 to 20 MPa, LHSV 0.5 to 3 h.^-1The hydrogen / oil ratio can be set according to the target hydrogenation rate among the conditions of 1000 to 5000 scfb. More preferably, the temperature is 320 to 380 ° C., the pressure is 5 to 7 MPa, and the LHSV is 0.5 to 3 h.^-1The hydrogen / oil ratio is 1000 to 3000 scfb.
[0011]
The heavy component containing three or more ring aromatic hydrocarbons separated by distillation needs to be hydrotreated at a higher pressure than the light component hydrogenation reaction.
As the catalyst, a catalyst in which cobalt and / or nickel and molybdenum are supported on a porous carrier mainly composed of alumina may be used. However, a catalyst having higher hydrogenation activity, for example, 85 to 99% by weight of alumina and zeolite. A catalyst including nickel and molybdenum supported on a support containing 1 to 15% by weight, or a catalyst supporting nickel and tungsten supported on 85 to 99% by weight of alumina or amorphous silica alumina and 1 to 20% by weight of zeolite is used. The hydrogenation rate can also be increased. In the case of these catalysts, the reaction conditions are as follows: temperature 320 to 380 ° C., pressure 3 to 20 MPa, LHSV 0.5 to 3 h.^-1The hydrogen / oil ratio of 1000 to 5000 scfb can be set according to the target hydrogenation rate, but it is necessary to set conditions for higher pressure and higher contact time than for light hydrogenation. Further preferable reaction conditions are as follows: temperature 320 to 360 ° C., pressure 7 to 15 MPa, LHSV 0.5 to 3 h.^-1The hydrogen / oil ratio is 2000 to 5000 scfb. Further, when the pressure is 10 MPa or more, a product oil having excellent hue can be obtained.
It is difficult to increase the size of a high-pressure device, and an increase in cost is inevitable for increasing the size. In the present invention, since aromatic hydrocarbons having three or more rings that are difficult to hydrogenate are separated and then hydrogenated under high pressure, the construction cost of the apparatus can be greatly reduced as compared with the case of treating all fractions with high pressure. .
[0012]
A catalyst in which a noble metal such as platinum, rhodium or palladium is supported on a porous carrier can also be used as a catalyst. When a noble metal catalyst is used, the reaction can be performed at a lower temperature and a lower pressure than these conditions. When these precious metal-supported catalysts are used, the reaction conditions for the preferred heavy component hydrotreating are as follows: temperature 200 to 300 ° C., pressure 1 to 5 MPa, LHSV 0.5 to 3 h.^-1The hydrogen / oil ratio is 1000 to 5000 scfb.
However, since the noble metal catalyst is subjected to sulfur poisoning, it is necessary to install a desulfurization reactor in the previous stage when processing raw material oil containing sulfur.
[0013]
Regarding the reaction temperature, the aromatic hydrogenation reaction is disadvantageous for the hydrogenation reaction from the parallel side of the reaction as the temperature is higher, so the reaction at an extremely high temperature should be avoided. Since the hydrogenation reaction for heavy components is difficult to proceed, a plurality of reactors arranged in parallel or in series may be used as necessary. In this case, hydrogen sulfide contained in the first stage reactor outlet hydrogen gas is separated by an amine absorber or the like, or a part of the hydrogen is extracted out of the system and replaced with new hydrogen instead. Decreasing the hydrogen concentration facilitates the hydrogenation reaction in the second stage. If the sulfur content can be reduced to 500 ppm or less in the first stage reaction, a noble metal catalyst such as platinum or palladium can be used in the second stage hydrogenation process. Regarding the light component, the hydrogenation process may be further advanced by installing a second reactor as necessary.
Precious metal catalysts are particularly effective for hydrogenation of heavy components, and it is not always necessary to use expensive precious metal catalysts for the hydrogenation reaction of all fractions. This is an important feature of the present invention. In the desulfurization treatment of the raw material oil, the entire fraction is desulfurized and then distilled to separate it into a soft component and a heavy component, and the heavy component may be treated with a precious metal catalyst, or only the heavy component is desulfurized after the distillation. Then, it may be hydrogenated with a noble metal catalyst.
[0014]
Thus, the heavy component sufficiently reduced in aromaticity can be mixed with the light component to produce a product. As for the heavy component, the light component may be mixed in the ratio when separated by distillation, and the distillation property of the product can be adjusted by changing the mixing ratio as necessary. It can also be commercialized by mixing with light oil produced by other desulfurization equipment. Of course, a lubricity improver, a cetane number improver, and a detergent can be blended as necessary when the product is produced by mixing.
[0015]
In this way, the light and heavy components that have been sufficiently reduced in aromatization can be mixed to produce a product. The light and heavy components may be mixed at a ratio when separated by distillation, and the distillation property of the product can be adjusted by changing the mixing ratio as necessary. It is also possible to produce a product by mixing with a light oil fraction treated in another hydrogenation apparatus. Of course, a lubricity improver, a cetane number improver, and a detergent can be blended as necessary when the product is produced by mixing.
[0016]
The amount of active metal supported on the catalyst used in the present invention may be the amount employed in a normal light oil hydrotreating catalyst. That is, assuming that the weight of the carrier is 100 parts by weight (weight including zeolite), Co or Ni is 1 to 10 parts by weight, preferably 3 to 6 parts by weight in terms of oxide, and Mo is 10 to 30 parts by weight in terms of oxide. Parts, preferably 15 to 25 parts by weight. If the amount of metal is small, the activity is insufficient and the deactivation rate of the catalyst is increased. On the other hand, if too much, the activity is saturated, which is uneconomical.
When using a catalyst in which nickel and molybdenum are supported on a support containing 1 to 15% by weight of zeolite, A zeolite, X zeolite, Y zeolite, L zeolite, MFI zeolite, mordenite, etc. can be used as the zeolite. . Among them, USY type zeolite obtained by dealumination of Y type zeolite to improve thermal stability is most preferable. These zeolites are ion-exchanged to develop Bronsted acid sites, but can be ion-exchanged with protons, alkaline earth metals, rare earth metals, and the like.
[0017]
Zeolite may be mixed with alumina gel, molded and fired, or may be adhered to the molded alumina support using a binder.
In order to improve the activity or the like of each hydrogen catalyst, a catalyst with a small amount of various reforming components added may be used. For example, when phosphorus is added, metal dispersion is improved and the Bronsted acid point is increased, so that hydrogenation activity may be improved. On the other hand, the addition of potassium or magnesium reduces the Bronsted acid point and suppresses the formation of thiols and sulfides, so that a sweet-nung effect can be expected in addition to the catalyst near the outlet of each reactor.
[0018]
The feedstock to which the present invention can be applied has a boiling range of 200 to 400 ° C. such as straight-run (decompressed) light oil, catalytic cracking (reduced pressure) light oil, thermal cracked (reduced pressure) light oil, etc., and bicyclic aromatic hydrocarbons and 3 A petroleum fraction containing aromatic hydrocarbons above the ring. The present invention is also effective for hydrotreating vacuum gas oil having a boiling point higher than 400 ° C.
The amount of aromatic hydrocarbon contained in the feedstock is not particularly limited, but is about 10% by weight in the case of ordinary straight-run gas oil. The amount of aromatic hydrocarbons in the product oil can be arbitrarily determined as required, and the required hydrogenation rate can be achieved by optimizing reaction conditions such as reaction temperature, pressure, liquid space velocity and the like.
Moreover, the raw material oil containing a sulfur compound and a nitrogen compound can also be used. In this case, hydrodesulfurization can be achieved simultaneously with aromatic hydrogenation to produce a low-sulfur product.
The light oil hydrogenated in the present invention can be used as a regular or premium diesel fuel for light oil vehicles. It can also be used by mixing with A heavy oil or the like.
[0019]
The reaction conditions of the present invention are as follows: temperature 320 to 380 ° C., pressure 3 to 20 MPa, LHSV 0.5 to 3 h.^-1The hydrogen / oil ratio can be set according to the target hydrogenation rate among the conditions of 1000 to 5000 scfb.
[0020]
The reactor used in the present invention may be any conventionally known reactor, for example, a fixed bed or a moving bed, and may be either a downflow type or an upflow type. Most suitable among these are fixed bed downflow reactors. Since this is a reactor mode conventionally used for hydrotreating petroleum fractions, the conventional apparatus can be used as it is. As the reactor, it is possible to use a reactor in which one reactor for light and heavy components is divided into a plurality of catalyst beds. The hydrotreating conditions of the present invention are so-called trickle beds in which a liquid and a gas coexist, and therefore it is desirable to install a distributor that uniformly disperses the liquid on each catalyst bed. Depending on the heat generation situation, quench hydrogen may be introduced at an optimal location to control the heat generation. In an actual apparatus, an extruded catalyst is used, and the catalyst is socked or densely charged into the reactor by conventional methods. After presulfiding the catalyst, the raw material oil heated together with hydrogen is passed through a reactor filled with the catalyst. A noble metal catalyst is used after being subjected to a reduction treatment without being presulfided. The spent catalyst may be used repeatedly by ordinary calcination regeneration treatment.
[0021]
【Example】
The invention is explained in more detail by means of examples.
Example 1
A Middle Eastern straight-run gas oil (boiling point 230-382 ° C.) was separated into 75% light by weight and 25% by weight heavy by using an atmospheric distillation apparatus with 20 theoretical plates and 340 ° C. as the cut temperature. When the proportion of aromatic hydrocarbons was examined using a mass spectrometer, the proportion of bicyclic aromatic hydrocarbons in the light component was 9.6% by weight, and the proportion of tricyclic aromatic hydrocarbons was 0.3% by weight. Met. The proportion of bicyclic aromatic hydrocarbons in the heavy component is 0.6% by weight, the proportion of tricyclic aromatic hydrocarbons is 4.3% by weight, and the proportion of tetracyclic aromatic hydrocarbons is 1.2% by weight.
As a first reactor, 5 parts by weight of cobalt (CoO equivalent) and 20 parts by weight of molybdenum (MoO) with respect to 100 parts by weight of γ-alumina support in a reaction tube having an inner diameter of 1 inch._Three300 ml of a catalyst carrying (conversion) was charged. This catalyst was used at 300 ° C., 5 MPa, LHSV1h using straight-run kerosene containing dimethyl disulfide (sulfur content 3% by weight).^-1After presulfiding for 4 hours under a hydrogen / oil ratio of 1000 scfb, the light component was heated to 350 ° C., pressure 5 MPa, LHSV 1 h.^-1Then, hydrogenation was performed by passing oil under conditions of a hydrogen / oil ratio of 1000 scfb. The ratio of bicyclic aromatic hydrocarbons in the product oil was 3.9% by weight, and the ratio of tricyclic aromatic hydrocarbons was 0.2% by weight.
Further, as a second reactor, 3 parts by weight of nickel (in terms of NiO) and 20 parts by weight of molybdenum are formed on a support containing 97% by weight of γ-alumina and 3% by weight of proton-exchanged USY zeolite in the upper layer of a 1-inch inner diameter reaction tube. (MoO_Three200 ml of the catalyst carrying the equivalent) was filled, and in the lower layer part, 5 parts by weight of cobalt (CoO equivalent) and 20 parts by weight of molybdenum (MoO equivalent) were added to the γ-alumina carrier._Three100 ml of the catalyst carrying the equivalent) was charged. This catalyst was used at 300 ° C., 5 MPa, LHSV1h using straight-run kerosene containing dimethyl disulfide (sulfur content 3% by weight).^-1, After presulfiding for 4 hours under the condition of hydrogen / oil ratio of 1000 scfb, the above-distilled heavy component was heated to 350 ° C., pressure 10 MPa, LHSV 1 h.^-1Then, hydrogenation was performed by passing oil under conditions of a hydrogen / oil ratio of 2000 scfb. The proportion of 2-ring aromatic hydrocarbons in the product oil is 0.0% by weight, the proportion of 3-ring aromatic hydrocarbons is 1.3% by weight, and the proportion of 4-ring aromatic hydrocarbons is 0.4% by weight. %Met. By mixing this heavy component and the aforementioned light component, the proportion of the bicyclic aromatic hydrocarbon is 2.9% by weight, the proportion of the tricyclic aromatic hydrocarbon is 0.5% by weight, A diesel oil with a proportion of aromatic hydrocarbons of 0.1% by weight and a color of Saybolt color (JISK-2580) was produced at +15.
[0022]
Example 2
Instead of the catalyst of Example 1, 5 parts by weight of nickel (in terms of NiO) and 20 parts by weight of molybdenum (MoO) are added to 100 parts by weight of the γ-alumina support in the first reactor and the second reactor._ThreeEach catalyst was loaded in an amount of 300 ml. This catalyst was presulfided in the same manner as in Example 1, and the light oil of Example 1 was desulfurized in the same condition as in Example 1 in the first reactor and the heavy part in the second reactor. The ratio of bicyclic aromatic hydrocarbons in the first reactor product oil is 3.6% by weight, and the ratio of tricyclic aromatic hydrocarbons is 0.1% by weight. The proportion of aromatic hydrocarbons was 0.0% by weight, the proportion of tricyclic aromatic hydrocarbons was 1.5% by weight, and the proportion of tetracyclic aromatic hydrocarbons was 0.5% by weight. By mixing this heavy component and the aforementioned light component, the ratio of the bicyclic aromatic hydrocarbon is 2.7% by weight and the ratio of the tricyclic aromatic hydrocarbon is 0.4% by weight. A diesel oil with a proportion of aromatic hydrocarbons of 0.1% by weight and a Saybolt color was used, producing +20 gas oil.
[0023]
Example 3
80% by volume of Middle Eastern straight oil (boiling point 224 to 368 ° C), 10% by volume of catalytic cracking gas oil (boiling point 212 to 345 ° C), and 10% by volume of direct decomposing light oil (boiling point 181 to 346 ° C) did. Was separated into 77% by volume of the light component and 23% by volume of the heavy component with a cutting temperature of 335 ° C. using an atmospheric distillation apparatus having 20 theoretical plates. The proportion of bicyclic aromatic hydrocarbons in the light component was 11.9% by weight, and the proportion of tricyclic aromatic hydrocarbons was 0.4% by weight. The proportion of bicyclic aromatic hydrocarbons in the heavy component is 0.7% by weight, the proportion of tricyclic aromatic hydrocarbons is 5.8% by weight, and the proportion of tetracyclic aromatic hydrocarbons is 1. 1% by weight.
The mixed light oil was loaded with the same amount of the same catalyst as in Example 1 in the first reactor, and the lighter portion in the second reactor charged with the same amount of the same catalyst as in Example 1. Hydrogenation was carried out under the same reaction conditions. The ratio of 2-ring aromatic hydrocarbons in the first reactor product oil is 4.2% by weight, and the ratio of tricyclic aromatic hydrocarbons is 0.2% by weight. The proportion of aromatic hydrocarbons was 0.0% by weight, the proportion of tricyclic aromatic hydrocarbons was 1.5% by weight, and the proportion of tetracyclic aromatic hydrocarbons was 0.4% by weight. By mixing this heavy component and the aforementioned light component, the proportion of the bicyclic aromatic hydrocarbon is 3.2% by weight and the proportion of the tricyclic aromatic hydrocarbon is 0.5% by weight. A diesel oil with a proportion of aromatic hydrocarbons of 0.1% by weight and a color of Saybolt color -2 was produced.
[0024]
Example 4
As a hydrodesulfurization reactor, 5 parts by weight of cobalt (CoO equivalent) and 20 parts by weight of molybdenum (MoO) with respect to 100 parts by weight of γ-alumina support in a reaction tube having an inner diameter of 1 inch._Three300 ml of a catalyst carrying (conversion) was charged. This catalyst was obtained by using straight-run kerosene (sulfur content: 3% by weight) containing dimethyl disulfide at a temperature of 300 ° C., a pressure of 5 MPa, and LHSV1h.^-1And presulfided for 4 hours under a hydrogen / oil ratio of 1000 scfb. Middle East straight gas oil (boiling point 230 to 382 ° C., sulfur content 1.5% by weight) used in Example 1 was temperature 350 ° C., pressure 5 MPa, LHSV 1h.^-1Then, hydrodesulfurization treatment was performed by passing oil under a hydrogen / oil ratio of 2000 scfb. The product oil was separated into 77% by volume of light component and 23% by volume of heavy component by cutting at 340 ° C. using an atmospheric distillation apparatus having 20 theoretical plates. When the proportion of aromatic hydrocarbons was examined with a mass spectrometer, the proportion of bicyclic aromatic hydrocarbons in the light component was 4.5% by weight, and the proportion of tricyclic aromatic hydrocarbons was 0.2% by weight. Met. In addition, the proportion of bicyclic aromatic hydrocarbons in the heavy component is 0.4% by weight, the proportion of tricyclic aromatic hydrocarbons is 2.3% by weight, and the proportion of tetracyclic aromatic hydrocarbons is 0.8% by weight. The sulfur content was 0.01% by weight for light and 0.04% by weight for heavy.
As a first reactor, in a reaction tube having an inner diameter of 1 inch, nickel 5 parts by weight (converted to Ni 2 O) and molybdenum 20 parts by weight (MoO) with respect to 100 parts by weight of γ-alumina support._Three300 ml of a catalyst carrying (conversion) was charged. This catalyst was obtained by using straight-run kerosene containing dimethyl disulfide (sulfur content: 3% by weight), temperature of 300 ° C., pressure of 5 MPa, LHSV1h.^-1, After presulfiding for 4 hours under the condition of hydrogen / oil ratio of 1000 scfb, the soft content is 350 ° C., pressure 5 MPa, LHSV 1 h.^-1Then, hydrogenation was performed by passing oil under conditions of a hydrogen / oil ratio of 1000 scfb. The ratio of bicyclic aromatic hydrocarbons in the product oil was 2.3% by weight, and the ratio of tricyclic aromatic hydrocarbons was 0.1% by weight. The sulfur content was 0.001% by weight. Further, as a second reactor, 200 ml of a catalyst supporting 0.5% by weight of platinum on a carrier containing 97% by weight of γ-alumina and 3% by weight of proton-exchanged USY zeolite was packed in a reaction tube having an inner diameter of 1 inch. After this catalyst was reduced with hydrogen at 300 ° C. for 4 hours, the heavy component separated by distillation was heated to 250 ° C., pressure 2 MPa, LHSV 1h.^-1And hydrogenation was performed by passing oil under conditions of a hydrogen / oil ratio of 2000 scfb. The ratio of bicyclic aromatic hydrocarbons in the product oil is 0.1% by weight, the ratio of tricyclic aromatic hydrocarbons is 1.0% by weight, and the ratio of tetracyclic aromatic hydrocarbons is 0.2% by weight. %Met. The sulfur content was 0.001% by weight.
By mixing this heavy component and the aforementioned light component, the ratio of the bicyclic aromatic hydrocarbon is 1.8% by weight and the ratio of the tricyclic aromatic hydrocarbon is 0.3% by weight. A light oil of +21 was produced with a proportion of aromatic hydrocarbons of 0.1% by weight and a color of Saybolt color (JISK-2580). The sulfur content was 0.002% by weight.
Comparative Example 1
In a reaction tube having an inner diameter of 1 inch, in a reaction tube, in a support containing 97% by weight of γ-alumina and 3% by weight of proton exchanged USY zeolite, 3 parts by weight of nickel (in terms of NiO) and 20 parts by weight of molybdenum (MoO)_Three300 ml of a catalyst carrying (conversion) was charged. Using straight-run kerosene containing dimethyl disulfide (sulfur content 3% by weight), 300 ° C, 5 MPa, LHSV 1h^-1After presulfiding for 4 hours under the condition of a hydrogen / oil ratio of 1000 scfb, the light oil used in the examples was heated to 350 ° C., pressure 5 MPa, LHSV 1 h.^-1Then, hydrogenation was performed by passing oil under conditions of a hydrogen / oil ratio of 1000 scfb. The proportion of bicyclic aromatic hydrocarbons in the product oil is 3.7% by weight, the proportion of tricyclic aromatic hydrocarbons is 1.0% by weight, and the proportion of tetracyclic aromatic hydrocarbons is 0.3%. The weight% and color were -15 in Saybolt color.
[0025]
Comparative Example 2
Into a reaction tube having an inner diameter of 1 inch, 5 parts by weight of cobalt (CoO equivalent) and 20 parts by weight of molybdenum (MoO) are added to 100 parts by weight of the γ-alumina support._Three600 ml of the catalyst carrying the equivalent) was charged. This catalyst was presulfided in the same manner as in Comparative Example 1, the diesel oil of Example 1 was passed through, and desulfurized under the same conditions as in Comparative Example 1. The proportion of bicyclic aromatic hydrocarbons in the product oil is 3.7% by weight, the proportion of tricyclic aromatic hydrocarbons is 1.2% by weight, and the proportion of tetracyclic aromatic hydrocarbons is 0.3%. Weight%, color was +15 in Saybolt color.
[0026]
Comparative Example 3
The mixed light oil of Example 3 was passed through the catalyst of Comparative Example 1 and hydrodesulfurized. The reaction conditions are the same as in Comparative Example 1. The ratio of 2-ring aromatic hydrocarbons in the product oil is 5.2% by weight, the ratio of 3-ring aromatic hydrocarbons is 2.5% by weight, and the ratio of 4-ring aromatic hydrocarbons is 0.5% by weight. %, The color was Seybolt color and -20 light oil was produced.
[0027]
【The invention's effect】
When hydrotreating a petroleum hydrocarbon fraction having a boiling point of 200 to 400 ° C. containing a polycyclic aromatic hydrocarbon, by adopting the present invention, the polycyclic aromatic content is low and there is almost no coloring. Excellent diesel oil can be produced efficiently.

Claims

When hydrotreating a fraction of petroleum hydrocarbons having a boiling point of 200 to 400 ° C. containing polycyclic aromatic hydrocarbons, 1 or more aromatic hydrocarbons having 3 or more rings contained in the light fraction separated by distillation are 1 The light fraction containing bicyclic aromatic hydrocarbons and the aromatic carbon of three or more rings by distillation so that the bicyclic aromatic hydrocarbons contained in the heavy component are 1% by weight or less. A method for hydrogenating polycyclic aromatic hydrocarbons, comprising separating into a heavy component containing hydrogen, hydrotreating the light component and the heavy component in independent steps, and then mixing the light component and the heavy component .

Comprising admixing a light fraction and a heavy fraction after treatment hydrogenated at light component and the heavy component or reactor either one of them were independent, the polycyclic aromatic hydrocarbon according to claim 1 Hydrogenation method.

The method for hydrogenating a polycyclic aromatic hydrocarbon according to claim 1 or 2 , wherein a cut temperature when separating a light component and a heavy component by distillation is 330 to 350 ° C.

In the hydrotreatment of light components, the catalyst support used is alumina, amorphous silica alumina, zeolite, or a mixture thereof, and a catalyst supporting cobalt and / or nickel and molybdenum on this support is used. The reaction conditions for hydrotreating the light component are a temperature of 320 to 380 ° C., a pressure of 5 to 7 MPa, an LHSV of 0.5 to 3 h ⁻¹ , and a hydrogen / oil ratio of 1000 to 3000 scfb. The method for hydrogenating polycyclic aromatic hydrocarbons according to Item .

After the sulfur content in the heavy component is set to 500 wtppm or less, a catalyst in which platinum, rhodium or palladium is supported on the porous support is used, and the reaction condition of the hydrogenation treatment for the heavy component is a temperature of 200 to 300. The method for hydrogenating a polycyclic aromatic hydrocarbon according to any one of claims 1 to 4 , wherein the pressure is ¹ ° C, 5 MPa, LHSV is 0.5 to 3 h ^-1 , and the hydrogen / oil ratio is 1000 to 5000 scfb.

The catalyst support used in the hydrotreatment of heavy components is alumina, amorphous silica alumina, zeolite, or a mixture thereof, and a catalyst having cobalt and / or nickel and molybdenum supported on this support is used. and the reaction conditions of the hydrogenation process for heavy components are temperature paths 320 to 360 ° C., a pressure 7~15MPa, LHSV0. 5~3h ^-1, a hydrogen / oil ratio 2000～5000Scfb, any claim 1-4 A method for hydrogenating a polycyclic aromatic hydrocarbon according to claim 1.